Fluidized bed process for upgrading diene-containing light olefins

ABSTRACT

An improved fluidized bed process for upgrading olefinic hydrocarbon feedstock by contacting the feedstock with acidic siliceous zeolite conversion catalyst particles at elevated temperature under exothermic conditions to produce heavier hydrocarbons including gasoline range hydrocarbons. The improvement comprises maintaining a turbulent fluidized bed of catalyst particles by flowing hydrocarbon-containing vapor upwardly through said bed at less than transport velocity; and introducing liquid olefinic feedstock comprising at least one C 4  -C 6  diene component into the fluidized catalyst bed in a lower portion thereof by rapidly atomizing and vaporizing the liquid feedstock, thus converting feedstock to heavier hydrocarbon without substantial thermal diene degradation thereof prior to contacting conversion catalyst particles in the fluidized bed. A predominantly liquid product is recovered containing C 4  -C 9  hydrocarbons rich in olefins and aromatics.

REFERENCE TO COPENDING APPLICATION

This application is a continuation-in-part of copending U.S. patentapplication Ser. No. 006,408 filed Jan. 23, 1987 now U.S. Pat. No.4,778,661, incorporated herein by reference.

BACKGROUND OF THE INVENTION

This invention relates to a catalytic process for upgrading olefinicstreams rich in dienes to heavier hydrocarbons rich in aliphatics andaromatics. In particular, it provides a continuous process foroligomerizing a feedstock containing monoalkenes and dienes to produceC₅ ⁺ hydrocarbons, such as liquid fuels, isobutane, aromatics and otheruseful products. Diene-containing liquids, such as thermal crackingliquids, are useful feedstocks herein.

Developments in zeolite catalysis and hydrocarbon conversion processeshave created interest in utilizing olefinic feedstocks for producing C₅⁺ gasoline, diesel fuel, etc. In addition to basic chemical reactionspromoted by ZSM-5 type zeolite catalysts, a number of discoveries havecontributed to the development of new industrial processes. These aresafe, environmentally acceptable processes for utilizing feedstocks thatcontain olefins. Conversion of C₂ -C₄ alkenes and alkanes to producearomatics-rich liquid hydrocarbon products were found by Cattanach (U.S.Pat. No. 3,760,024) and Yan, et al (U.S. Pat. No. 3,845,150) to beeffective processes using the ZSM-5 type zeolite catalysts. In U.S. Pat.Nos. 3,960,978 and 4,021,502, Plank, Rosinski and Givens discloseconversion of C₂ -C₅ olefins, alone or in admixture with paraffiniccomponents, into higher hydrocarbons over crystalline zeolites havingcontrolled acidity. Garwood, et al have also contributed to theunderstanding of catalytic olefin upgrading techniques and improvedprocesses as in U.S. Pat. Nos. 4,150,062, 4,211,640 and 4,227,992. Theabove-identified disclosures are incorporated herein by reference.

Conversion of olefins, especially propene and butenes, over HZSM-5 iseffective at moderately elevated temperatures and pressures. Theconversion products are sought as liquid fuels, especially the C₅ ⁺aliphatic and aromatic hydrocarbons and C₄ hydrocarbons, in particulariso-butane. Product distribution for liquid hydrocarbons can be variedby controlling process conditions, such as temperature, pressure andspace velocity. Gasoline (C₅ -C₁₀) is readily formed at elevatedtemperature [(e.g., up to about 700° C.)] and moderate pressure fromambient to about 5500 kPa, preferably about 200 to 2900 kPa. Olefinicgasoline can be produced in good yield and may be recovered as a productor fed to a low severity, high pressure reactor system for furtherconversion to heavier distillate-range products. Distillate modeoperation can be employed to maximize production of C₁₀ ⁺ aliphatics byreacting the lower and intermediate olefins at high pressure andmoderate temperature. Operating details for typical "MOGD"oligomerization units are disclosed in U.S. Pat. Nos. 4,456,779;4,497,968 (Owen, et al.) and 4,433,185 (Tabak), incorporated herein byreference. At moderate temperature and relatively high pressure, theconversion conditions favor distillate-range product having a normalboiling point of at least 165° C. (330° F.).

Many feedstocks of commercial interest, such as thermal crackingbyproduct, etc., contain both mono-olefins and diolefins (e.g. C₂ -C₆mono-alkenes and C₄ ⁺ dienes) along with C₁ -C₁₀ light aliphatics, and aminor amount of aromatics. Gaseous and liquid streams containing dienesare typically produced in thermal cracking operations. One commonexample is pyrolysis gasoline, which is produced as ethene (ethylene)byproduct. Such diene-containing streams are often difficult to processdue to poor thermal stability and the tendency of dienes to form cokeand gum deposits. This complicates preheating of such streams into thehigh temperatures required of most catalytic upgrading processes. Priorattempts to upgrade such materials have pretreated the feedstock tohydrogenate the diolefin selectively, as in U.S. Pat. No. 4,052,477(Ireland et al). The present invention is concerned with providing asafe and low cost technique for catalytically converting diene-richstreams to high value C₄ ⁺ products rich in aromatics.

It has been found that diene-containing olefinic light hydrocarbons canbe upgraded directly to liquid hydrocarbons rich in C₅ ⁺ aliphatics andaromatics by catalytic conversion in a turbulent fluidized bed of solidacid zeolite catalyst under high severity reaction conditions withoutdeleterious effects from the diolefin components. This technique isparticularly useful for upgrading C₄ ⁺ liquid pyrolysis products, whichmay contain minor amounts of ethene, propene, C₂ -C₄ paraffins andhydrogen produced in cracking petroleum fractions, such as naphtha,ethane or the like. By upgrading the complex olefinic by-product,gasoline yield of cracking units can be significantly increased.Accordingly, it is a primary object of the present invention to providea novel technique for upgrading diene-rich hydrocarbon streams.

SUMMARY OF THE INVENTION

A process has been found for continuous conversion of diene-containingolefinic feedstock to heavier hydrocarbon products wherein the feedstockis contacted at elevated temperature with a fluidized bed of acidicsiliceous zeolite catalyst under high severity conversion conditions.The improvement herein comprises the steps of maintaining a turbulentfluidized bed of catalyst particles by flowing hydrocarbon-containingvapor upwardly through the bed at less than transport velocity and byintroducing liquid olefinic feedstock comprising at least one C₄ -C₆diene component into the fluidized catalyst bed in a lower bed portionby rapidly atomizing and vaporizing the liquid feedstock, thusconverting feedstock to heavier hydrocarbon without substantial thermaldiene degradation thereof prior to contacting conversion catalystparticles in the fluidized bed. This technique prevents fouling ofconduits, furnaces and other upstream equipment.

In a preferred embodiment of the invention, the olefinic feed comprisesabout 5 to 90 wt. % total C₄ ⁺ mono-olefin and 5 to 50 wt. % conjugatedunsaturated dienes, and the fluidized bed is maintained at an averagetemperature of about 315 to 510° C. The preferred catalyst comprisesmedium pore shape selective metallosilicate.

In a turbulent fluidized catalyst bed the conversion reactions areconducted in a vertical reactor column by passing feedstock gas upwardlythrough the reaction zone at a velocity greater than dense bedtransition velocity and less than transport velocity for the averagecatalyst particle. A continuous process is operated by withdrawing aportion of coked catalyst from the reaction zone, oxidativelyregenerating the withdrawn catalyst and returning regenerated catalystto the reaction zone at a rate to control catalyst activity and reactionseverity whereby propane:propene molar ratio in the hydrocarbon productis maintained at about 0.2:1 to 200:l under conditions of reactionseverity to effect feedstock conversion. A thermodynamically balancedmixture of exothermic alkenes and endothermic alkanes can be convertedwithout significant recycle and/or diluent. However, a supplementalfeedstream or recycle stream such as C₃ ⁻ hydrocarbons can be introducedinto the reactor bed. Such a recycle stream can increase C₅ ⁺ aliphaticand aromatic yields, while lowering catalyst makeup requirements.

THE DRAWINGS

FIG. 1 is a schematic view of a fluidized bed reactor system accordingto the present invention;

FIG. 2 is a vertical cross section view of a liquid-gas feed nozzlewhich is employed to introduce low-temperature diene feed into thereactor bed;

FIG. 3 is an aging plot showing the effect of adding about 1 to 5% of adiene (butadiene) to a C₂ -C₄ olefinic feed; and

FIG. 4 is a process flow sheet depicting system operations, including afractionation section for product recovery and recycle fluid handling.

DESCRIPTION OF PREFERRED EMBODIMENTS Description of Catalysts

Recent developments in zeolite technology have provided a group ofmedium pore siliceous materials having similar pore geometry. Mostprominent among these intermediate pore size zeolites is ZSM-5, which isusually synthesized with Bronsted acid active sites by incorporating atetrahedrally coordinated metal, such as Al, Ga, B, Fe or mixturesthereof, within the zeolitic framework. These medium pore zeolites arefavored for acid catalysis; however, the advantages of ZSM-5 structuresmay be utilized by employing highly siliceous materials or cystallinemetallosilicate having one or more tetrahedral species having varyingdegrees of acidity. ZSM-5 crystalline structure is readily recognized byits X-ray diffraction pattern, which is described in U.S. Pat. No.3,702,866 (Argauer, et al.), incorporated by reference.

The oligomerization catalysts preferred for use herein include themedium pore (i.e., about 5-7A) shape-selective crystallinealuminosilicate zeolites having a silica-to-alumina ratio of at least12, a constraint index of about 1 to 12 and acid cracking activity(alpha value) of about 10-250, preferably about 10 to 80 based on totalcatalyst weight. In the fluidized bed reactor the coked catalyst mayhave an apparent activity (alpha value) of about 10 to 80 under theprocess conditions to achieve the required degree of reaction severity.Representative of the ZSM-5 type medium pore shape selective zeolitesare ZSM-5, ZSM-11, ZSM-12, ZSM-22, ZSM-23, ZSM-35, ZSM-38, and ZSM-48.Aluminosilicate ZSM-5 is disclosed in U.S. Pat. No. 3,702,886 and U.S.Pat. No. Re. 29,948. Other suitable zeolites are disclosed in U.S. Pat.Nos. 3,709,979; 3,832,449; 4,076,979; 3,832,449; 4,076,842; 4,016,245and 4,046,839; 4,414,423; 4,417,086; 4,517,396 and 4,542,251. Thedisclosures of these patents are incorporated herein by reference. Whilesuitable zeolites having a coordinated metal oxide to silica molar ratioof 20:1 to 200:1 or higher may be used, it is advantageous to employ astandard ZSM-5 having a silica alumina molar ratio of about 25:1 to70:1, suitably modified if desired to adjust acidity and aromatizationcharacteristics. A typical zeolite catalyst component having Bronstedacid sites may consist essentially of aluminosilicate ZSM-5 zeolite with5 to 95 wt. % silica and/or alumina binder.

These siliceous zeolites may be employed in their acid forms ionexchanged or impregnated with one or more suitable metals, such as Ga,Pd, Zn, Ni, Co and/or other metals of Periodic Groups III to VIII. Thezeolite may include a hydrogenation-dehydrogenation component (sometimesreferred to as a hydrogenation component) which is generally one or moremetals of group IB, IIB, IIIB, VA, VIA or VIIIA of the Periodic Table(IUPAC), especially aromatization metals, such as Ga, Pd, etc. Usefulhydrogenation components include the noble metals of Group VIIIA,especially platinum, but other noble metals, such as palladium, gold,silver, rhenium or rhodium, may also be used. Base metal hydrogenationcomponents may also be used, especially nickel, cobalt, molybdenum,tungsten, copper or zinc. The catalyst materials may include two or morecatalytic components, such as a metallic oligomerization component (eg,ionic Ni⁺², and a shape-selective medium pore acidic oligomerizationcatalyst, such as ZSM-5 zeolite) which components may be present inadmixture or combined in a unitary bifunctional solid particle. It ispossible to utilize an ethene dimerization metal or oligomerizationagent to effectively convert feedstock ethene in a continuous reactionzone. Certain of the ZSM-5 type medium pore shape selective catalystsare sometimes known as pentasils. In addition to the preferredaluminosilicates, the borosilicate, ferrosilicate and "silicalite"materials may be employed.

ZSM-5 type pentasil zeolites are particularly useful in the processbecause of their regenerability, long life and stability under theextreme conditions of operation. Usually the zeolite crystals have acrystal size from about 0.01 to 2 microns or more. In order to obtainthe desired particle size for fluidization in the turbulent regime, thezeolite catalyst crystals are bound with a suitable inorganic oxide,such as silica, alumina, etc. to provide a zeolite concentration ofabout 5 to 95 wt. %. It is advantageous to employ a standard ZSM-5having a silica:alumina molar ratio of 25:1 or greater in a once-throughfluidized bed unit to convert 60 to 100 percent, preferably at least 75wt. %, of the monoalkenes and dienes in the feedstock. In thedescription of preferred embodiments a 25% H-ZSM-5 catalyst calcinedwith 75% silica-alumina matrix binder is employed unless otherwisestated.

Particle size distribution can be a significant factor in achievingoverall homogeneity in turbulent regime fluidization. It is desired tooperate the process with particles that will mix well throughout thebed. Large particles having a particle size greater than 250 micronsshould be avoided, and it is advantageous to employ a particle sizerange consisting essentially of 1 to 150 microns. Average particle sizeis usually about 20 to 100 microns, preferably 40 to 80 microns.Particle distribution may be enhanced by having a mixture of larger andsmaller particles within the operative range, and it is particularlydesirable to have a significant amount of fines. Close control ofdistribution can be maintained to keep about 10 to 25 wt. % of the totalcatalyst in the reaction zone in the size range less than 32 microns.This size range of fluidizable particles is classified as Geldart GroupA. The fluidization regime is controlled to assure operation between thetransition velocity and transport velocity, and these fluidizationconditions are substantially different from those found in non-turbulentdense beds or transport beds.

Process Operation

In this description, metric units and parts by weight are employedunless otherwise stated.

Suitable olefinic feedstocks comprises C₄ -C₆ alkenes includingconjugated dienes such as 1,3-butadiene, pentadiene, isomers,hexadienes, cyclic dienes, or similar C₄ ⁺ aliphatic liquid hydrocarbonshaving diethylenic conjugated unsaturation. Aromatics coproduced withthe liquid olefinic components may be cofed or separated by solventextraction prior to conversion of the diene-rich feedstock.Non-deleterious components, such as paraffins and inert gases, may bepresent. A particularly useful feedstock is a liquid by-product ofpyrolysis or thermal cracking units containing typically 40-95 wt. % C₄-C₆ total mono-olefins and di-olefins, including about 5-60 wt. % diene,along with varying amounts of C₃ -C₈ paraffins, aromatics and inerts.Specific examples are given in Table 1 below. The process may betolerant of a wide range of lower alkanes, from 0 to 95%. Preferredpyrolysis feedstocks contain more than 50 wt. % C₄ -C₆ lower aliphatichydrocarbons, and contain sufficient olefins to provide an olefinicpartial pressure of at least 50 kPa. Under the high severity reactionconditions employed in the present invention, lower alkanes may bepartially converted to heavier hydrocarbons.

The desired products are C₅ to C₉ hydrocarbons, which will ordinarilycomprise at least 50% of the recovered product, preferrable 80% or more.While olefins may be a predominant fraction of the C₅ ⁺ reactioneffluent; it is often desired to upgrade the feedstock to high octanegasoline containing aromatics, preferrably at least 5% C₆ -C₈ aromatics(BTX).

The reaction severity conditions can be controlled to optimize yield ofC₄ -C₉ hydrocarbons. It is understood that aromatics and light paraffinproduction is promoted by those zeolite catalysts having a highconcentration of Bronsted acid reaction sites. Accordingly, an importantcriterion is selecting and maintaining catalyst inventory to provideeither fresh or regenerated catalyst having the desired properties.Typically, acid cracking activity (alpha value) can be maintained fromhigh activity values greater than 100 to significantly lower valuesunder steady state operation by controlling catalyst deactivation andregeneration rates to provide an apparent average alpha value below 100,preferably about 15 to 80.

Reaction temperatures and contact time are also significant factors inthe reaction severity, and the process parameters are followed to give asubstantially steady state condition wherein the reaction severity index(R.I.) is maintained within the limits which yield a desired weightratio of propane to propene. While this index may vary from about 0.2 to200, it is preferred to operate the steady state fluidized bed unit tohold the R.I. below about 50, with optimum operation at 0.7 to 2 in thesubstantial absence of added propane. While reaction severity isadvantageously determined by the weight ratio of propane:propene in thegaseous phase, it may also be approximated by the analogous ratios ofbutanes:butenes, pentanes;pentenes, or the average of total reactoreffluent alkanes:aklenes in the C3-C5 range. Accordingly, thesealternative expressions may be a more accurate measure of reactionseverity conditions when propane is added to the feedstock. The optimalvalue will depend upon the exact catalyst composition, feedstock andreaction conditions; however, the typical diene-rich feed mixtures usedin the examples herein and additional olefinic feeds can be optimallyupgraded to the desired aliphatic-rich gasoline by keeping the R.I. atabout 1.

Upgrading of olefins by such hydrogen contributors in fluidized bedcracking and oligomerization units is taught by Owen et al in U.S. Pat.No. 4,090,949. This technique is particularly useful for operation witha pyrolysis cracking unit to increase overall production of liquidproduct.

The use of fluidized bed catalysis permits the conversion system to beoperated at low pressure drop, which in an economically practicaloperation can provide a maximum operating pressure only 50 to 200 kPaabove atmospheric pressure. Another important advantage is the closetemperature control that is made possible by turbulent regime operation,wherein the uniformity of conversion temperature can be maintainedwithin close tolerances, often less than 5° C. Except for a small zoneadjacent the bottom gas inlet, the midpoint measurement isrepresentative of the entire bed, due to the thorough mixing achieved.

Referring now to FIG. 1, a reactor vessel 2 is shown provided with heatexchange tube means 4. There may be several separate heat exchange steamgenerating tube bundles so that temperature control can be separatelyexercised over the fluid catalyst bed. The bottoms of the tubes arespaced above a feed distributor grid 8 sufficiently to be free of jetaction by the charged gas passing through the small diameter holes inthe grid 8. although depicted without baffles, the vertical reactionzone can contain open end tubes above the grid for maintaining hydraulicconstraints, as disclosed in U.S. Pat. No. 4,251,484 (Daviduk andHaddad). Optionally, a variety of horizontal baffles may be added tolimit axial mixing in the reactor. Heat released from the reaction canbe controlled by adjusting feed temperature in a known manner. A largeportion of reaction heat can be removed by feeding cold liquid into thereactor at a temperature at least 200° C. below average bed temperature.In the reactor configuration shown the heat exchanger tubes can functionas dummy tubes to limit mixing in the reactor.

The system provides for withdrawing catalyst from above grid 8 byconduit means 10. This flow line can be provided with control valvemeans 12 for passage to catalyst regeneration in vessel 13, where cokedcatalyst particles are oxidatively regenerated in contact with air orother regeneration gas at high temperature. The oxidatively regeneratedcatalyst is then passed to the reactor fluid bed of catalyst by conduitmeans 14 and flow control valve 16. The regenerated catalyst is chargedto the catalyst bed sufficiently below the upper interface to achievegood mixing in the fluid bed. Since the flow of regenerated catalystpassed to the reactor can be relatively small, hot regenerated catalystdoes not ordinarily upset the temperature constraints of the reactoroperations in a significant amount.

Initial fluidization is achieved by forcing a lift gas upwardly througha catalyst, a light aliphatic C₄ ⁻ gas, with or without diluent orrecycle, may be charged through inlet port 20A at a bottom portion ofthe reactor in open communication with chamber 24 beneath grid 8.Pressurized feedstock is introduced above reactant distributor grid 8via supply conduit 21, pump 22 and distributor conduit 23 to one or morespray nozzle means, described and depicted in FIG. 2. The liquid isdispersed into the bed of catalyst thereabove at a velocity sufficientto form a generally upwardly flowing suspension of atomized liquidreactant with the catalyst particles and lift gas.

Advantageously, the liquid diene-containing reactant feed is injectedinto the catalyst bed by atomizing the pressurized liquid feedstream toform readily dispersible liquid particles having an average size of 300microns or less. This contributes to rapid vaporization of the liquid atprocess pressure. Exothermic conversion provides sufficient heat tovaporize the liquid quickly, thus avoiding deleterious liquid phasereactions of the diene components, which tend to form carbonaceousdeposits such as heavy coke, gums, etc.

A plurality of sequentially connected cyclone separator means 30, 32 and34 provided with diplegs 36, 38 and 40 respectively are positioned in anupper portion of the reactor vessel comprising dispersed catalyst phase28.

The product effluent separated from catalyst particles in the cycloneseparating system then passes to a plenum chamber 42 before withdrawalvia conduit 46, operatively connect with effluent separation system 50.The product effluent is cooled and separated to recover C₅ ⁺ liquidhydrocarbons, gaseous recycle or offgas, along with any byproduct wateror catalyst fines carried over. A portion of the light gas effluentfraction may be recycled by compressing to form a motive gas for theliquid feed or via recycle conduit 20B for use as lift gas. Therecovered hydrocarbon product comprising C₅ ⁺ olefins and/or aromatics,paraffins and naphthenes is thereafter processed as required to providea desired gasoline or higher boiling product.

Under optimized process conditions the turbulent bed has a superficialvapor velocity of about 0.3 to 2 meters per second (m/sec). At highervelocities entrainment of fine particles may become excessive and beyond10 m/sec the entire bed may be transported out of the reaction zone. Atlower velocities, the formation of large bubbles or gas voids can bedetrimental to conversion. Even fine particles cannot be maintainedeffectively in a turbulent bed below about 0.1 m/sec.

A convenient measure of turbulent fluidization is the bed density. Atypical turbulent bed has an operating density of about 100 to 500kg/m³, preferrably about 300 to 500, measured at the bottom of thereaction zone, becoming less dense toward the top of the reaction zonedue to pressure drop and particle size differentiation. This density isgenerally between the catalyst concentration employed in dense beds andthe dispersed transport systems. Pressure differential between twovertically spaced points in the reactor column can be measured to obtainthe average bed density at such portion of the reaction zone. Forinstance, in a fluidized bed system employing ZSM-5 particles having aclean apparent density of 1.06 gm/cc and packed density of 0.85, anaverage fluidized bed density of about 300 to 500 kg/m³ is satisfactory.

By virtue of the turbulence experienced in the turbulent regime,gas-solid contact in the catalytic reactor is improved, providingsubstantially complete conversion, enhanced selectivity and temperatureuniformity. One main advantage of this technique is the inherent controlof bubble size and characteristic bubble lifetime. Bubbles of thegaseous reaction mixture are small, random and short-lived, thusresulting in good contact between the gaseous reactants and the solidcatalyst particles.

A significant difference between the process of this invention andconversion processes of the prior art is that operation in the turbulentfluidization regime is optimized to produce high octane C₅ ⁺ liquid ingood yield. The weight hourly space velocity and uniform contactprovides a close control of contact time between vapor and solid phases,typically about 3 to 25 seconds. Another advantage of operating in sucha mode is the control of bubble size and life span, thus avoiding largescale gas by-passing in the reactor. The process of the presentinvention does not rely on internal baffles in the reactor for thepurpose of bubble size control such as the baffles which are employed inthe prior art dense bed processes discussed above.

As the superficial gas velocity is increased in the dense bed,eventually slugging conditions occur and with a further increase in thesuperficial gas velocity the slug flow breaks down into a turbulentregime. The transition velocity at which this turbulent regime occursappears to decrease with particle size. The turbulent regime extendsfrom the transition velocity to the so-called transport velocity, asdescribed by Avidan, et al in U.S. Pat. No. 4,547,616 and by Tabak, etal. in U.S. Pat. No. 4,579,999, incorporated herein by reference. As thetransport velocity is approached, there is a sharp increase in the rateof particle carryover, and in the absence of solid recycle, the bedcould empty quickly.

Several useful parameters contribute to fluidization in the turbulentregime in accordance with the process of the present invention. Whenemploying a ZSM-5 type zeolite catalyst in fine powder form such acatalyst should comprise the zeolite suitably bound or impregnated on asuitable support with a solid density (weight of a representativeindividual particle divided by its apparent "outside" volume) in therange from 0.6-2 g/cc, preferably 0.9-1.6 g/cc. The catalyst particlescan be in a wide range of particle sizes up to about 250 microns, withan average particle size between about 20 and 100 microns, preferably inthe range of 10-150 microns and with the average particle size between40 and 80 microns. When these solid particles are placed in a fluidizedbed where the superficial fluid velocity is 0.3-2, operation in theturbulent regime is obtained. The velocity specified here is for anoperation at a total reactor pressure of about 100 to 300 kPa. Thoseskilled in the art will appreciate that at higher pressures, a lower gasvelocity may be employed to ensure operation in the turbulentfluidization regime.

The reactor can assume any technically feasible configuration, butseveral important criteria should be considered. The bed of catalyst inthe reactor can be at least about 5-20 meters in height, preferablyabout 7 meters. Fine particles may be included in the bed, especiallydue to attrition, and the fines may be entrained in the product gasstream. A typical turbulent bed may have a catalyst carryover rate up toabout 1.5 times the reaction zone inventory per hour. If the fraction offines becomes large, a portion of the carryover can be removed from thesystem and replaced by larger particles. It is feasible to have a fineparticle separator, such as a cyclone disposed within the reactor shellto recover catalyst carryover and return this fraction continuously tothe bottom of the reaction zone for recirculation at a rate of about onecatalyst inventory per hour. Optionally, fine particles carried from thereactor vessel entrained with effluent gas can be recovered by a highoperating temperature sintered metal filter.

This process can be used with any process stream which containssufficient liquid olefins and dienes and is substantially free ofdeleterious oxygenates and sulfur compounds. Experimental runs areperformed using a ZSM-5 catalyst to demonstrate the inventive process.The fluidized bed unit can be operated over a wide range of processvariables and catalyst activity.

Reactor Operation

A typical single pass reactor unit employs a temperature-controlledcatalyst zone with indirect heat exchange and/or adjustable gas quench,whereby the reaction exotherm can be carefully controlled to preventexcessive temperature above the usual operating range of about 315° C.to 510° C., preferably at average reactor temperature of 340° C. to 430°C. Energy conservation in the system may utilize at least a portion ofthe reactor exotherm heat value by exchanging hot reactor effluent withfeedstock and/or recycle streams. Optional heat exchangers may recoverheat from the effluent stream prior to fractionation. It is preferred tooperate the olefin conversion reactors at moderate pressure of about 100to 3000 kPa (atmospheric to about 400 psig).

The weight hourly space velocity (WHSV, based on total olefins in thefresh feedstock is about 0.1-5 WHSV. Typical product fractionationsystems are described in U.S. Pat. Nos. 4,456,779 and 4,504,693 (Owen,et al).

In order to prevent premature non-catalytic reaction of the dienes, itis desirable to maintain reactant liquid feedstream temperature belowabout 180° C. (350° F.) until injection into the fluidized bed.Appropriate thermal insulation or quenching of the feedstream to theinjection point can largely prevent gum and coke formation in the liquidphase prior to catalysis.

Atomization of the pressurized liquid reactant feedstream can beachieved by known techniques, such as liquid spray nozzles, motive gas,ultra sonics, etc. A suitable nozzle is shown in FIG. 2, wherein aconcentric feed liquid projection device 100 is depicted in verticalcross section view. Pressurized liquid flows through a supply conduit123. The nozzle is mounted onto the vessel internal structure by screwcap means 130 or similar attachment means. A motive fluid supplied underpressure through conduit 126 drives the pressurized liquid flowing fromthe nozzle orifice 140 for injection into the reaction vessel atsufficient velocity to induce a fine vertically directed spray ofatomized liquid having an average particle size up to about 300,preferably about 50 microns. The number and arrangement of nozzles willdepend upon the cross sectional area of the fluidized bed andfluidization characteristics of the gas-solid-liquid mixture. Theatomized stream from a pressurized nozzle can be made to effectpenetration into the bed at a depth and/or lateral radius of a meter ormore. The mixture fluid may be an inert material, nitrogen, loweraliphatic gas, stream, etc.

Thermal insulation of the liquid diene-containing feedstream from thehot reaction medium in the reaction vessel can be achieved by applyingto the liquid feed conduit a layer of thermal insulation, such as aceramic shield or the like. Jacketed conduits with heat adsorbing fluidmay also be suitable.

EXAMPLE 1

In the present example a C₄ ⁺ liquid stream is converted toaromatics-rich gasoline in the fluidized bed reactor employing acidZSM-5 powder catalyst having a fresh alpha value of about 80 at anaverage conversion temperature about 425° C. (800° F.) and totalpressure of about 275 kPa (25 psig). The liquid pyrolysis gasolinefeedstock contains about 22 wt. % C₄ ⁺ mono-alkenes, 27% C₄ ⁺ dienes(mainly 1,3-butadiene), 49% C₄ ⁺ paraffins, 2% aromatics and naphthenes,and less than 1% C₃ ⁻ aliphatics. Following initial heating andfluidization of the powdered catalyst with a heated lift gas (e.g. C₂ ⁻hydrocarbon), the feedstream is preheated and maintained below 180° C.prior to injection into the bed. After achieving steady state operationat a reaction severity index (R.I.) of about 1, the effluent conversionproduct (less any lift gas components) comprises 82 wt. % C₅ ⁺ liquidgasoline having a research octane rating of 94 (RON). The totalaromatics content is 18 wt. %, including 1% benzene (B), 5% toluene (T),6% xylenes (X) and ethyl benzene, 4% C₉ aromatics isomers and 10% C₁₀isomers, mainly durene. The predominant nonaromatic fraction (65%)contains mainly mono-olefins, paraffins and naphthenes, and the lightgas C₄ ⁻ fraction is 17% of the conversion product.

Typical olefinic pyrolysis byproduct streams are shown in Table 1.

                  TABLE 1                                                         ______________________________________                                        Example of Diene-Rich Feestock (ethane cracker byproduct)                     Component        Vol. %                                                       ______________________________________                                        C.sub.3.sup.-    1.0                                                          i-butene         0.08                                                         1,3-butadiene    0.51                                                         t.2, butene      0.1                                                          c.2, butene      0.15                                                         1,2 butadiene    0.14                                                         3m 1 butene      0.45                                                         isopentane       5.44                                                         1,4 pentadiene   0.6                                                          1-pentene        0.63                                                         n-pentane        1.92                                                         isoprene         2.3                                                          c,2, pentene     0.35                                                         2m 2 butene      0.45                                                         t,1,3, pentadiene                                                                              1.5                                                          c,1,3, pentadiene                                                                              1.0                                                          cyclopentadiene  13.7                                                         cyclopentene     1.7                                                          2,3 d.m. butane  1.7                                                          3m pentene       0.85                                                         hexane           0.95                                                         unknown C.sub.6  1.04                                                         cyclohexane      3.06                                                         benzene          34.4                                                         unknown C.sub.8  3.47                                                         Toluene          10.1                                                         vinyleydohexene  0.19                                                         ethylbenzene     1.29                                                         xylene           1.01                                                         styrene          0.3                                                          unknown C.sub.9.sup.+                                                                          6.9                                                          ______________________________________                                    

The above diene-rich stream example contains C₆ ⁺ aromatic hydrocarbonswhich can be separated before feeding to the reactor. Typical ranges ofdiene-rich pyrolysis gasoline streams comprised of mainly C₄ -C₆hydrocarbons are:

    ______________________________________                                                      Vol. %                                                          ______________________________________                                               Dienes   5-60                                                                 Mono-alkenes                                                                           5-30                                                                 Aromatics                                                                              1-5*                                                                 Alkanes  20-60                                                                Naphthenes                                                                             1-5                                                           ______________________________________                                    

EXAMPLES 2-4

A series of continuous olefin conversion runs are conducted usingH-ZSM-5 (65%) catalyst having an alpha value of about 175 at thebeginning of the aging runs made under oligomerization conditionswithout regeneration to upgrade mixtures of ethene, propene andbutadiene and to determine the effects of diene concentration oncatalyst aging. The control feedstock (Example 2) is compared withdiene-containing feeds in Table 1.

                  TABLE 2                                                         ______________________________________                                                     Example 2                                                                             Example 3 Example 4                                      ______________________________________                                        Ethene         0         0.7       1.8                                        Propene        26.8      28.1      22.9                                       Butenes        35.7      31.9      31.7                                       1,3 Butadiene  0 (control)                                                                             0.8       5.1                                        Alkanes (C.sub.4.sup.-)                                                                      37.5      38.5      38.5                                       Recycle (mol/mol olefin)                                                                     2.5:1     2.5:1     2.5:1                                      ______________________________________                                    

The conversion unit is a single bed isothermal reactor employingparticulate zeolite having a crystal size less than 0.5 microns,together with 35% alumina binder. The continuous runs are conducted atabout 6600 kPa and weight hourly space velocity (WHSV) of about 0.8parts olefin feed per part by weight of catalyst per hour. Theconversion runs are started at 205° C. (400° F.) and the temperature isincreased to compensate for coke deposition, while maintaining totalolefin conversion of at least 80%, preferably over 90%. Results of theaging studies are plotted in FIG. 3, with all conversion rates beingnormalized to 80% to 330° F.+ product for comparison purposes.Selectivity of the conversion product to heavier hydrocarbons is shownin Table 3.

                  TABLE 3                                                         ______________________________________                                                     Example 2                                                                             Example 3 Example 4                                      ______________________________________                                        Total Liquid Product,                                                         50% pt, °C. (°F.)                                                              261/(501) 259/(498) 244/(472)                                  Distillate Species (As Cut)                                                   5 wt. %, °C. (°F.)                                                             232/(434) 250/(483) 297(477)                                   95 wt. % °C. (°F.)                                                             369/(697) 383/(722) 379(715)                                   Gravity, °API                                                                         44.3      41.2      38.9                                       Aniline Point  177       184       172                                        ______________________________________                                    

While the aromatics product content of the control runs averaged about2-5%, the 5.1% butadiene feed (Example 4) is upgraded to an aromaticscontent of 15.5 wt. %, more than 3 times the diene input. The averageparaffin content is less than 14% and the liquid dominant product is70% + olefins and naphthenes.

These results indicate butadiene, at levels of 1 wt. % or less, do notcause significantly increased catalyst aging or lower productselectivity. Typical FCC C₃ /C₄ olefins from a depropanizer feed streamcontain 0.3-0.6 wt. % butadiene which is less than the 0.8 wt. %butadiene concentration that was used in this study. Even at the 5.1 wt.% butadiene level, though catalyst again was increased, productselectivity to heavier hydrocarbons remained relatively high.

The flexibility of the fluid bed operating parameters for controllingthe reactor temperature under exothermic reaction conditions allows aneasy adjustment for achieving the optimal yield structure.

In order to effect fluidization of the catalyst at the bottom of thereactor prior to injection of the liquid feed stream, a lift gas may beemployed. This can be an inert diluent or recyled light gas, such asmethane, ethane, ethene, propane, etc. Recycle of C₃ ⁻ lighthydrocarbons may also be desirable under certain circumstances, forinstance with unreacted aliphatics which require further conversion orfor dilution of highly exothermic feedstocks. The thermodynamic balanceof exothermic olefin oligomerization and endothermic paraffin reactionscan have significant impact on the reaction severity conditions.

In FIG. 4, the feedstock stream 201 is injected into reactor vessel 210containing the fluidized bed of catalyst, along with fluidizing gasstream 214 and recycle streams 216, 218. Reactor effluent is cooled inheat exchanger 220 and partially separated in a series of phaseseparation drums 222 (high temperature separator) and 224 (lowtemperature separator). A light gas stream may be recovered from LTS224, pressurized in compressor 128, and recycled via conduit 214 tocomprise at least a portion of the lift gas. Condensed liquid from theseparators 222, 224 is fed to a first fractionation tower 230, alongwith a portion of the LTS overhead vapor. Overhead vapor is furtherfractionated by second light hydrocarbon fractionation tower 232 fromwhich offgas stream 234, C₃ -C₄ LPG etc. are recovered. Liquid bottomsstream 236, rich in C₅ and optionally some LPG alkanes, may be recycledvia 218 for further conversion via conduit 218 to reactor 210. The C6+liquid from the tower 230 is passed to a product splitter fractionationunit 240 for recovery of C₆ -C₈ hydrocarbons rich in BTX aromaticscomponents. The C₉ ⁺ aliphatic and aromatic components may be recoveredas a heavy gasoline product stream; however, it is advantageous torecycle this stream via 216 for further conversion to increase the netaromatic product.

The use of a fluid-bed reactor in this process offers several advantagesover a fixed-bed reactor. Due to continuous catalyst regeneration,fluid-bed reactor operation will not be adversely affected by oxygenate,sulfur and/or nitrogen containing contaminants present in the pyrrolysisbyproduct.

While the invention has be shown by describing preferred embodiments ofthe process, there is no intent to limit the inventive concept, exceptas set forth in the following claims.

We claim:
 1. A fluidized bed catalytic process for upgrading diene-richliquid olefinic feedstock comprising at least 1 wt. % diene and a totalC₄ -C₆ olefin content of about 5 to 90 wt. % to produce hydrocarbonsrich in C5+ aliphatics and aromatics, comprising the steps ofmaintaininga fluidized bed of zeolite catalyst particles in a turbulent reactor bedat a temperature of about 220° to 510° C., said catalyst having anapparent particle density of about 0.9 to 1.6 g/cm³ and a size range ofabout 1 to 150 microns, and average catalyst particle size of about 20to 100 microns; feeding the olefinic feedstock as a pressurized liquidinto a bottom portion of the fluidized bed; dispersing and vaporizingthe liquid feedstock in the fluidized bed; passing vaporized hotfeedstock upwardly through the fluidized catalyst bed under turbulentflow conditions; maintaining turbulent fluidized bed conditions throughthe reactor bed between transition velocity and transport velocity at asuperficial fluid velocity of about 0.3 to 2 meters per second; andrecovering hydrocarbon product containing a major amount of C₄ ⁺hydrocarbons and containing C₃ -C₅ alkanes and aklenes in the weightratio of about 0.2:1 to 200:1.
 2. A fluidized bed process according toclaim 1 wherein the fluidized bed density is about 100 to 500 kg/m³,measured at the bottom of the bed, and has bed height of at least 7meters; and wherein the catalyst comprises a siliceous metallosilicateacid zeolite having the structure of ZSM-5 zeolite containing about 10to 25 weight percent of fine particles having a particle size less than32 microns.
 3. The process of claim 1 wherein the feedstock consistsessentially of C₄ -C₆ aliphatic liquid comprising about 5 to 50 weightpercent dienes and is substantially free of oxygenates and organosulfurcompounds and wherein said liquid is fed in an atomized stream of finelydivided liquid particles.
 4. A process for continuous conversion ofdiene-containing hydrocarbon feedstock to heavier hydrocarbon productswherein the feedstock is contacted at elevated temperature with afluidized bed of zeolite catalyst under conversion conditions,comprising the steps of:maintaining the fluidized catalyst bed in avertical reactor column having a turbulent reaction zone by passingvapor upwardly through the reaction zone at a velocity greater thandense bed transition velocity to a turbulent regime and less thantransport velocity for the average catalyst particle; injecting anatomized stream of liquid feedstock into the reaction zone; andwithdrawing a portion of coked catalyst from the reaction zone,oxidatively regenerating the withdrawn catalyst and returningregenerated catalyst to the reaction zone at a rate to control catalystactivity whereby propane:propene molar ratio in the hydrocarbon productis maintained at about 0.2:1 to 50:1 in the substantial absence of addedpropane under conditions of reaction severity to effect feedstockconversion.
 5. The process of claim 4 wherein the superficial feedstockvapor velocity is about 0.3-2 m/sec; the reaction temperature is about315 to 510° C.; the weight hourly feedstock space velocity (based ontotal olefin) is about 0.1 to 5; propane:propene weight ratio about0.7:1 to 2:1; and the average fluidized bed density measured at thereaction zone bottom is about 300 to 500 kg/m³.
 6. The process of claim4 wherein the catalyst consists essentially of a medium pore pentasilzeolite having an apparent alpha value of about 10 to 250, and averageparticle size of about 20 to 100 microns, including at least 10 weightpercent fine particles having a particle size less than 32 microns. 7.The process of claim 6 wherein the catalyst particles comprise about 5to 90 weight percent ZSM-5 zeolite having a crystal size of about 0.01-2microns.
 8. The process of claim 4 wherein said feedstock consistsessentially of C₄ -C₆ hydrocarbon liquid containing about 5 to 50 wt. %C₄ -C₆ conjugated unsaturated dienes and substantially free ofdeleterious oxygenates and sulfur compounds and without significantrecycle.
 9. The process of claim 8 wherein the olefin partial pressureis at least 50 kPa.
 10. In the process for upgrading olefinichydrocarbon feedstock by contacting the feedstock with acidic siliceouszeolite conversion catalyst particles at elevated temperature underexothermic conditions to produce heavier hydrocarbons including gasolinerange hydrocarbons; the improvement which comprises:maintaining aturbulent fluidized bed of catalyst particles by flowinghydrocarbon-containing vapor upwardly through said bed at less thantransport velocity; and introducing liquid olefinic feedstock comprisingat least one C₄ -C₆ diene component into the fluidized catalyst bed in alower portion thereof by rapidly atomizing and vaporizing the liquidfeedstock, thus converting feedstock to heavier hydrocarbon withoutsubstantial thermal diene degradation thereof prior to containingconversion catalyst particles in the fluidized bed.
 11. The process ofclaim 10 wherein the olefinic feed comprises about 5 to 90 wt. % C₄ ⁺mono-alkene and 5 to 50 wt. % conjugated unsaturated dienes, wherein thefluidized bed is maintained at an average temperature of about 315 to510° C. and the catalyst comprises medium pore shape selectivemetallosilicate.
 12. The process of claim 11 wherein said feedstockconsists essentially of liquid pyrolysis gasoline byproduct ofhydrocarbon pyrolysis, wherein the catalyst comprises zeolite having thecrystalline structure of ZSM-5, wherein the super ficial vapor velocitythrough the fluidized bed is about 0.2 to 3 meters/sec and the liquidfeedstock is introduced at super atmosphere pressure and at atemperature of at least 200° C. below process temperature of the bed.13. The process of claim 12 wherein the liquid feedstock is atomized toan average particle size not greater than about 300 microns.
 14. Theprocess of claim 10 wherein C₃ ⁻ hydrocarbon gas is added to liquid feedto provide motive gas for atomization.
 15. The process of claim 10wherein a portion of light hydrocarbon gas recovered from reactoreffluent is recycled to the reactor as a fluidizing lift gas and isdistributed to the bottom of the reactor through a grid below the liquidfeed.
 16. The process of claim 10 wherein the catalyst is modified witha hydrogenation-dehydrogenation metal component to increase aromaticsproduction.
 17. A continuous process for catalytic conversion ofdiolefin-containing hydrocarbon feedstock to heavier hydrocarbonproducts wherein the feedstock is converted at elevated temperature,comprising the steps of:containing a fluidized bed of zeolite catalystin a vertical reactor; maintaining the fluidized catalyst bed in avertical reactor column having a turbulent reaction zone by passingvapor upwardly through the reaction zone at a velocity greater thandense bed transition velocity to a turbulent regime and less thantransport velocity for the average catalyst particle; injecting anatomized stream of liquid feedstock into the reaction zone; withdrawinga portion of coked catalyst from the reaction zone; and reactivating thewithdrawn catalyst and returning regenerated catalyst to the reactionzone at a rate to control catalyst activity to maintain conditions ofreaction severity to effect feedstock conversion.
 18. The process ofclaim 17 including atomizing liquid with at least one spray nozzleadapted to received a liquid feedstream and motive gas and having anupwardly directed spray orifice for projecting atomized liquid into thefluidized bed from a bottom portion thereof.
 19. A continuous processfor converting lower olefinic hydrocarbon feedstock to aromatics-richproduct by contacting the feedstock in a fluidized bed reaction zonewith acidic siliceous zeolite conversion catalyst particles at elevatedtemperature under exothermic conditions to produce heavier hydrocarbonsincluding gasoline range hydrocarbons comprising:maintaining a turbulentfluidized bed of catalyst particles by flowing hydrocarbon-containingvapor upwardly through said bed at less than transport velocity; andintroducing lower olefinic feedstock comprising at least one C₄ -C₆diene component into the fluidized catalyst bed in a lower portionthereof to convert feedstock to heavier hydrocarbon rich in aromatichydrocarbon without substantial thermal diene degradation thereof priorto contacting conversion catalyst particles in the fluidized bed;separating from fluidized bed reaction effluent (a) a firstaromatics-rich liquid stream consisting essentially of C₆ -C₈hydrocarbons, (b) a second light hydrocarbon stream comprising C₅ ⁻aliphatic hydrocarbons, and (c) a third heavy C₉ ⁺ hydrocarbons stream;and recycling at least a portion of the second C₅ ⁻ aliphatics-richstream or at least a portion of the third C₉ ⁺ heavy hydrocarbon streamto the fluidized bed reaction zone.
 20. The process of claim 19 whereinliquid feedstock is injected at super-atmosphere pressure as an atomizedliquid.
 21. The process of claim 20 wherein the liquid feedstockcomprises pyrolysis gasoline liquid consisting essentially of 5-60volume % dienes, 5-30% mono-alkenes, 1-5% aromatics, 20-60% alkanes and1-5% naphthenes.
 22. The process of claim 20 wherein the liquidfeedstock comprises pyrolysis gasoline liquid containing about 5-60volume % dienes, 5-30% mono-alkenes, and 0-60% aromatics.
 23. Theprocess of claim 19 wherein reaction severity conditions are maintainedby controlling catalyst acidity, reaction temperature and reactantcontact time to produce reaction effluent containing propane and propenein the ratio of about 0.2:1 to 200:1.